This invention relates to catalytic cracking of hydrocarbons. It more particularly refers to improvements in the endothermic catalytic cracking of petroleum fractions and alternating exothermic catalyst regeneration.
Endothermic catalytic cracking of hydrocarbons, particularly petroleum fractions, to lower molecular weight desirable products is well known. This process is practiced industrially in a cycling mode wherein hydrocarbon feedstock is contacted with hot, active, solid particulate catalyst without added hydrogen at rather low pressures of up to about 50 psig and temperatures sufficient to support the desired cracking. As the hydrocarbon feed is cracked to lower molecular weight, more valuable and desirable products, "coke" is deposited on the catalyst particles. The coked catalyst is disengaged from the hydrocarbon products, which are then resolved and separated into appropriate components. The coked catalyst particles, now cooled from the endothermic cracking and disengaged from the hydrocarbon products, are then contacted with an oxygen containing gas whereupon coke is burned off the particles to regenerate their catalytic activity. During regeneration, the catalyst particles absorb the major portion of the heat generated by the combustion of coke, i.e. they are "reflexively" heated, with consequent increase of catalyst temperature. The heated, regenerated catalyst particles are then contacted with additional hydrocarbon feed and the cycle repeats itself.
A flue gas comprising carbon oxides is produced during regeneration. In conventional operation this flue gas contains substantial quantities of carbon monoxide. The carbon monoxide is either vented to the atmosphere with the rest of the flue gas or is in some way burned to carbon dioxide, in an incinerator or a CO boiler or the like.
It has recently become desirable to decrease the content of carbon monoxide in the regenerator flue gas for at least two reasons. In the first place, CO combustion is extremely exothermic and in view of the increasing cost of energy, burning CO in the regenerator increases the heat efficiency of the reflexive endothermic catalytic cracking system. In the second place, since carbon monoxide is an air pollutant, more and more stringent controls are being placed upon its venting into the environment. It is therefore clearly desirable to provide means for burning carbon monoxide within a reflexive hydrocarbon catalytic cracking system. This has been attempted in the past and is being attempted at present by means of increasing the temperature and air input to the regenerator so as to support thermal combustion of carbon monoxide in the regenerator. This technique has been difficult to commercialize and to operate successfully in a smooth, steady state manner.
In the past attempts have been made, in fact it has sometimes been commercial practice, to employ special catalysts for this process which contain a cracking component and a component for catalyzing the oxidation of carbon monoxide. The CO oxidation components used in the past have been metals of the transition element group and/or of the iron group. In particular, manganese, cobalt and especially chromium have been used for this purpose.
Two major variants for endothermically cracking hydrocarbons are fluid catalytic cracking (FCC) and moving bed catalytic cracking. In both of these processes as commercially practiced, the feed hydrocarbon and the catalyst are passed through a "reactor"; are disengaged; the catalyst is regenerated with cocurrent and/or countercurrent air; and the regenerated reflexively heated catalyst recontacted with more feed to start the cycle again. These two processes differ substantially in the size of the catalyst particles utilized in each and also in the engineering of materials contact and transfer which is at least partially a function of the catalyst size.
In fluid catalytic cracking (FCC), the catalyst is a fine powder of about 10 to 200 microns, preferably about 70 micron, size. This fine powder is generally propelled upwardly through a riser reaction zone suspended in and thoroughly mixed with hydrocarbon feed. The coked catalyst particles are separated from the cracked hydrocarbon products, and after purging are transferred into the regenerator where coke is burned to reactivate the catalyst. Regenerated catalyst generally flows downward from the regenerator to the base of the riser.
One typical example of industrially practiced moving bed hydrocarbon catalytic cracking is known as thermofor catalytic cracking (TCC). In this process the catalyst is in the shape of beads or pellets having an average particle size of about 1/64 to 1/4 inch, preferably about 1/2 inch. Active, hot catalyst beads progress downwardly cocurrent with a hydrocarbon charge stock through a cracking reaction zone. In this zone hydrocarbon feed is endothermically cracked to lower molecular weight hydrocarbons while coke is deposited on the catalyst. At the lower end of the reaction zone the hydrocarbon products are separated from the coked catalyst, and recovered. The coked catalyst is then passed downwardly to a regeneration zone, into which air is fed such that part of the air passes upwardly countercurrent to the coked catalyst and part of the air passes downwardly cocurrent with partially regenerated catalyst. Two flue gases comprising carbon oxides are produced. Regenerated catalyst is disengaged from the flue gas and is then lifted, pneumatically or mechanically, back up to the top of the reaction zone.
The catalysts used in endothermic catalytic nonhydrogenative cracking are to be distinguished from catalysts used in exothermic catalytic hydrocracking. Operating conditions also to be distinguished. While the catalytic cracking processes to which this invention is directed operate at low pressures near atmospheric and in the absence of added hydrogen, hydrocracking is operated with added hydrogen at high pressures of up to about 1,000 to 3,000 psig. Further, non-hydrogenative catalytic cracking is a reflexive process with catalyst cycling between cracking and regeneration (coke burn off) over a very short period of time, seconds or minutes. In hydrocracking, on the other hand, the catalyst remains in cracking service for an extended period of time, months, between regeneration (coke burn off). Another important difference is in the product. Non-hydrogenative catalytic cracking produces a highly unsaturated product with substantial quantities of olefins and aromatics, and a high octane gasoline fraction. Hydrocracking, in contrast produces an essentially olefin-free product with a relatively low octane gasoline.
This invention is not directed to hydrocracking nor is it within the scope of this invention to use hydrocracking catalysts in the process hereof. Hydrocracking catalysts have an acidic cracking component, which may be a crystalline aluminosilicate zeolite, amorphous silica alumina, clays or the like, and a very strong hydrogenation/dehydrogenation component. Strong hydrogenation/dehydrogenation components are illustrated by metals such as molybdenum, chromium and vanadium, and group VIII metals such as cobalt, nickel and palladium. These are used in relatively large proportion, certainly large enough to support heavy hydrogenation of the charge stock under the conditions of hydrocracking. To the contrary, strong hydrogenation/dehydrogenation metals are neither required nor desired as components of non-hydrogenative catalytic cracking. In fact, it is usual for some metals, such as nickel and vanadium, to deposit out on the catalyst from the charge stock during non-hydrogenative cracking. These are considered to be catalyst poisons in this process and therefore to be avoided or at least minimized. Their detrimental effect in non-hydrogenative catalytic cracking is to increase the coke and light gas, including hydrogen, produced in the cracking reaction and therefore to reduce the yield of desired liquid products, particularly gasoline.
FIG. 1 and the sectional element thereof shown in FIG. 2 are representative of a commercial fluid catalytic cracking unit. Referring now to FIG. 1, a hydrocarbon feed 2 such as a gas oil boiling from about 600.degree. F. up to 1,000.degree. F. is passed after preheating thereof to the bottom portion of riser 4 for admixture with hot regenerated catalyst introduced by standpipe 6 provided with flow control valve 8. A suspension of catalyst in hydrocarbon vapors at a temperature of at least about 950.degree. F. but more usually at least 1,000.degree. F. is thus formed in the lower portion of riser 4 for flow upwardly therethrough under hydrocarbon conversion conditions. The suspension initially formed in the riser may be retained during flow through the riser for a hydrocarbon residence time in the range of 1 to 10 seconds.
The hydrocarbon vapor-catalyst suspension formed in the riser reactor is passed upwardly through riser 4 under hydrocarbon conversion conditions of at least 900.degree. F. and more usually at least 1,000.degree. F. before discharge into one or more cyclonic separation zones about the riser discharge, represented by cyclone separator 14. There may be a plurality of such cyclone separator combinations comprising first and second cyclonic separation means attached to or spaced apart from the riser discharge for separating catalyst particles from hydrocarbon vapors. Separated hydrocarbon vapors are passed from separator 14 to a plenum chamber 16 for withdrawal therefrom by conduit 18. These hydrocarbon vapors together with gasiform material separated by stripping gas as defined below are passed by conduit 18 to fractionation equipment not shown. Catalyst separated from hydrocarbon vapors in the cyclonic separation means is passed by diplegs represented by dipleg 20 to a dense fluid bed of separated catalyst 22 retained about an upper portion of riser conversion zone 4. Catalyst bed 22 is maintained as a downwardly moving fluid bed of catalyst counter-current to rising gasiform material. The catalyst passes downwardly through a stripping zone 24 immediately therebelow and counter-current to rising stripping gas introduced to a lower portion thereof by conduit 26. Baffles 28 are provided in the stripping zone to improve the stripping operation.
The catalyst is maintained in stripping zone 24 for a period of time sufficient to effect a higher temperature desorption of feed deposited compounds which are then carried overhead by the stripping gas. The stripping gas with desorbed hydrocarbons pass through one or more cyclonic separating means 32 wherein entrained catalyst fines are separated and returned to the catalyst bed 22 by dipleg 34. The hydrocarbon conversion zone comprising rise 4 may terminate in an upper enlarged portion of the catalyst collecting vessel with the commonly known bird cage discharge device or an open end "T" connection may be fastened to the riser discharge which is not directly connected to the cyclonic catalyst separation means. The cyclonic separation means may be spaced apart from the riser discharge so that an initial catalyst separation is effected by a change in velocity and direction of the discharged suspension so that vapors less encumbered with catalyst fines may then pass through one or more cyclonic separation means before passing to a product separation step. In any of these arrangements, gasiform materials comprising stripping gas hydrocarbon vapors and desorbed sulfur compounds are passed from the cyclonic separation means represented by separator 32 to a plenum chamber 16 for removal with hydrocarbon products of the cracking operation by conduit 18. Gasiform material comprising hydrocarbon vapors is passed by conduit 18 to a product fractionation step not shown. Hot stripped catalyst at an elevated temperature is withdrawn from a lower portion of the stripping zone by conduit 36 for transfer to a fluid bed of catalyst being regenerated in a catalyst regeneration zone. Flow control valve 38 is provided in transfer conduit 36.
This type of catalyst regeneration operation is referred to as a swirl type of catalyst regeneration due to the fact that the catalyst bed tends to rotate or circumferentially circulate about the vessel's vertical axis and this motion is promoted by the tangential spent catalyst inlet to the circulating catalyst bed. Thus, the tangentially introduced catalyst at an elevated temperature is further mixed with hot regenerated catalyst or catalyst undergoing regeneration at an elevated temperature and is caused to move in a circular or swirl pattern about the regenerator's vertical axis as it also moves generally downward to a catalyst withdrawal funnel 40 (sometimes called the "bathtub") adjacent the regeneration gas distributor grid. In this catalyst regeneration environment, it has been found that the regeneration gases comprising flue gas products of carbonaceous material combustion tend to move generally vertically upwardly through the generally horizontally moving circulating catalyst to cyclone separators positioned above the bed of catalyst in any given vertical segment. As shown by FIG. 2, the catalyst tangentially introduced to the regenerator by conduit 36 causes the catalyst to circulate in a clockwise direction in this specific embodiment. As the bed of catalyst continues its circular motion some catalyst particles move from an upper portion of the mass of catalyst particles suspended in regeneration gas downwardly therethrough to a catalyst withdrawal funnel 40 in a segment of the vessel adjacent to the catalyst inlet segment. In the regeneration zone 42 housing a mass of the circulating suspended catalyst particles 44 in upflowing oxygen containing regeneration gas introduced to the lower portion thereof by conduit distributor means 46, the density of the mass of suspended catalyst particles may be varied by the volume of regeneration gas used in any given segment or segments of the distributor grid. Generally speaking, the circulating suspended mass of catalyst particles 44 undergoing regeneration with oxygen containing gas to remove carbonaceous deposits by burning will be retained as a suspended mass of swirling catalyst particles varying in density in the direction of catalyst flow and a much less dense phase of suspended catalyst particles 48 will exist thereabove to an upper portion of the regeneration zone. Under carefully selected relatively low regeneration gas velocity conditions, a rather distinct line of demarcation may be made to exist between a dense fluid bed of suspended catalyst particles and a more dispersed suspended phase (dilute phase) of catalyst thereabove. However, as the regeneration gas velocity conditions are increased there is less of a demarcation line and the suspended catalyst passes through regions of catalyst particle density generally less than about 30 lbs. per cu. ft. A lower catalyst bed density of at least 20 lb/cu. ft. is preferred.
A segmented regeneration gas distributor grid 50 positioned in the lower cross-sectional area of the regeneration vessel 42 is provided as shown in FIG. 1 and is adapted to control the flow of regeneration gas passed to any given vertical segment of the catalyst bed thereabove. In this arrangement, it has been found that even with the generally horizontally circulating mass of catalyst, the flow of regeneration gas is generally vertically upwardly through the mass of catalyst particles so that regeneration gas introduced to the catalyst bed by any given grid segment or portion thereof may be controlled by grid openings made available and the air flow rate thereto. Thus, oxygen containing combustion gases after contact with catalyst in the regeneration zone are separated from entrained catalyst particles by the cyclonic means provided and vertically spaced thereabove. The cyclone combinations diagrammatically represented in FIG. 1 are intended to correspond to that represented in FIG. 2. Catalyst particles separated from the flue gases passing through the cyclones are turned to the mass of catalyst therebelow by the plurality of provided catalyst diplegs.
As mentioned above, regenerated catalyst withdrawn by funnel 40 is conveyed by standpipe 6 to the hydrocarbon conversion riser 4.
The regenerator system shown in FIGS. 1 and 2 is usually designed for producing a flue gas that contains a substantial concentration of carbon monoxide along with carbon dioxide. In fact, a typical CO.sub.2 /CO ratio is about 1.2.
As noted above, there has recently been a marked increase in the desire to reduce carbon monoxide emissions from the regenerator of a reflexive non-hydrogenative catalytic cracking process. Prior proposed solutions, of increasing the temperature of the regenerator sufficient to thermally burn CO, or of incorporating chromium or iron with the cracking catalyst to support catalytic CO combustion, have not accomplished a sufficient reduction in CO emissions or, when this reduction has approached sufficiency, it has been at the expense of a great detriment to the operation and product distribution of the cracking reaction side of this process. In addition to the fact that increased production of coke on the cracking side throws this entire reflexive system into heat imbalance, the increased production of light gas unduly strains the capacity of the compressors and the entire gas plant, that is the series of separation operation in which the C4 gasiform part of the product is resolved into its component parts.
It is therefore an important object of this invention to provide a novel means of reducing carbon monoxide emissions from a reflexive, non-hydrogenative catalytic cracking process.
It is another object of this invention to provide a novel catalyst for such a process.
It is another object of this invention to provide an improved process for the fluid catalytic cracking of gas oils in the absence of added hydrogen:
Other and additional objects of this invention will become apparent from a consideration of this entire specification including the claims hereof.
Various references which have been uncovered may bear upon the subject matter of this application. The following list is not suggested to be exhaustive or all inclusive. It does, however, identify those presently known references which, in the opinion of counsel, seem to bear upon this subject matter:
U.S. Pat. No. 3,660,310 Kluksdahl; PA1 U.S. Pat. No. 3,788,977 Dolbear; PA1 U.S. Pat. No. 3,364,136 Chen et al.; PA1 U.S. Pat. No. 2,436,927 Kassel PA1 U.S. Pat. No. 3,226,339 Frilette et al.